Auto thermal reforming (ATR) catalytic systems

ABSTRACT

An autothermal reforming catalytic structure for generating hydrogen gas from liquid hydrocarbons, steam and an oxygen source. The autothermal reforming catalytic structure includes a support structure and nanosized mixed metal oxide particles dispersed homogenously throughout the support structure.

CROSS-REFERENCE TO RELATED PATENT APPLICATIONS

This application is a continuation of and claims priority to U.S. patentapplication Ser. No. 13/083,899 titled “AUTO THERMAL REFORMING (ATR)CATALYTIC STRUCTURES” and filed on Apr 11, 2011 now U.S. Pat. No.9,745,191.

TECHNICAL FIELD OF THE INVENTION

The present invention relates to the formulation of autothermalreforming (ATR) catalytic structures that are useful for hydrogengeneration from liquid fuel hydrocarbons.

BACKGROUND OF THE INVENTION

With stricter environmental regulations on transportation fuels, coupledwith higher level of heavier crude oil processing, demand for hydrogenin refineries is expected to increase in order to meet the need for morehydrotreating and hydrocracking processes. Today, about 95% of hydrogenis supplied by large-scale steam reforming of natural gas, which employsreforming, water-gas-shift (WGS), and pressure adsorption processes insuccessive steps. Efficiency, however, suffers for smaller scaleoperations. Therefore, hydrogen production at a smaller scale fromfossil fuels necessitates further development to meet the requirementsof purity, economics and versatility for a variety of emergingapplications such as the use of fuel cell powered vehicles.

During the last decade, there have been growing concerns about theshortage of energy supply in the world. Fuel cell technologies that usehydrogen may offer potentially non-polluting and efficient fuel fortoday's rising energy demands. While this type of technology is indevelopment for renewable sources of hydrogen, hydrogen is typicallycommercially produced through conversion of hydrogen rich energycarriers.

Therefore, cost effective methods for the production of hydrogen fromfossil fuels are becoming increasingly important, and more particularly,methods for the production of low cost, pure hydrogen, which is freefrom carbon monoxide are needed.

Hydrocarbon steam reforming (SR), partial oxidation (PDX) andautothermal reforming (ATR) are important known processes for hydrogenproduction. Steam reforming of high molecular weight hydrocarbons hasbeen practiced for over 40 years in locations where natural gas is notavailable. ATR of higher hydrocarbons, however, is preferable to SRbecause it combines a highly endothermic SR process with an exothermicPDX process. This means that ATR is a more thermally stable process thanSR, and can be driven to thermo-neutral conditions. In addition, thestart up of the ATR is more rapid than SR.

Sulfur deposition on many commercial catalysts tends to induce cokeaccumulation, and because commercial fuels contain sulfur that cannot becompletely removed (due to the sulfur being present in the aromatic ringof the fuels), use of these types of commercial catalysts typicallyresults in unwanted coking and early deactivation of the catalysts.Conversely, catalysts having noble metals, particularly Rh-basedcatalysts, can tolerate sulfur and prevent coke deposition, therebymaking them more active in the reforming of hydrocarbons to synthesisgas. A significant drawback of noble metal based catalysts, however, istheir significantly greater cost, as compared with non-noble metal basedcatalysts. For example, the price of rhodium has risen almost 10 foldover the past 3 years from an average of about USD $16.77 per gram in2007 to about $149.71 per gram in 2010.

Steam reforming of methane is typically used as a form of catalyticreaction for the production of hydrogen. Conventional catalytic systemsutilizing either nickel or noble metal catalysts for the steam reformingof methane require primary reaction temperatures of about 700° C. andabove, followed by rather extensive and expensive purification processesto provide a hydrogen product with reasonable purity (i.e., greater than95% by volume) that can be used as a feed stock for many commonprocesses.

Reforming of hydrocarbons other than methane, notably oil products, canbe achieved with reactions similar to steam reforming of natural gas.The successive breaking of the C—C terminal bonds of higher hydrocarbons(i.e., hydrocarbons having at least 2 carbon atoms present) withsuitable catalysts is more difficult than for methane, however, due todifferences in reaction rates and increased propensity for thermalcracking (also called pyrolysis). To avoid the issues related todiffering reaction rates and thermal cracking, carbon stripping isusually done in a separate pre-reformer, making this option of producinghydrogen more complex and more expensive than natural gas reforming.

Hydrogen can also be produced from non-volatile hydrocarbons, such asheavy oils, by gasification or partial oxidation. Gasification processesutilize steam at temperatures above about 600° C. to produce hydrogenwhilst carbon is oxidized to CO₂. Gasification, however, is usually noteconomical compared to steam reforming and partial oxidation processes.By comparison, partial oxidation is a considerably more rapid processthan steam reforming. In partial oxidation, the hydrocarbon feed ispartially combusted with oxygen in the presence of steam at atemperature of between about 1300° C. and 1500° C. Pressure has littleeffect on the reaction rate of the process and is usually performed atpressures of about 2 to 4 MPa to permit the use of more compactequipment and compression cost reduction. When air is used as the oxygensource, nitrogen must be removed from the resulting hydrogen gasproduct, typically requiring a separate stage following the oxidationreactor. Partial oxidation is thus more suitable for small scaleconversion, such as in a motor vehicle that is equipped with fuel cells.The process can be stopped and started, as required for on-boardreformation, and when in progress, it can provide elevated temperaturesthat may start steam reforming along with the oxidation processes. Thisis called autothermal reforming and involves all the reactions mentionedso far.

The water-gas-shift (WGS) reaction is an alternative hydrogen productiontechnology frequently used following a primary catalytic reactionutilized to remove carbon monoxide impurities and increase overallhydrogen yield. The WGS reaction is mildly exothermic and thus isthermodynamically favored at lower temperatures. The kinetics of thereaction, however, are superior at higher temperatures. Therefore, it iscommon practice to first cool the reformate product from the reformer ina heat exchanger to a temperature between about 350° C. and 500° C., andthen conduct the reaction over a suitable WGS catalyst. The resultingreformate is then cooled once again to a temperature between about 200°C. and 250° C., and reacted on a low temperature designated WGScatalyst. Due to these several conversion and heat exchanging stepsinvolved, however, the process is economically expensive and highlyinefficient.

Pressure Swing Adsorption (PSA) is a well-known established method toseparate hydrogen from a stream containing impurities. PSA employsmultiple beds, usually two or more, of solid adsorbents to separate theimpure hydrogen stream into a very pure (99.9%) hydrogen product streamand a tail gas stream that includes the impurities and a fraction of thehydrogen produced. As an example, synthesis gas (H₂ and CO) may beintroduced into one bed where the impurities, rather than the hydrogen,are adsorbed onto the adsorbents. Ideally, just before complete loadingis achieved, this adsorbent bed is switched offline and a secondadsorbent bed is placed on line. The pressure on the loaded bed issubsequently reduced, which liberates the adsorbed impurities (in thiscase predominantly CO₂) at low pressure. A percentage of the inlethydrogen, typically approximately 15 percent, is lost in the tail gas. Asignificant disadvantage of the PSA is that low tail gas pressureessentially limits the system to a single WGS stage. Limiting a hydrogenseparation system to a single WGS stage thus decreases the amount of COconversion as well as the total amount of hydrogen recovery. PSA is alsoundesirable, as compared to the use of membranes, in part due to themechanical complexity of the PSA, which leads to higher capital andoperating expenditures and potentially increased downtime.

Hydrogen is produced and removed through hydrogen permeable metalmembranes, such as palladium or palladium alloys. Metallic membranes,particularly palladium or palladium alloys, however, are very expensive,sensitive to sulfur compounds, and difficult to co-sinter with or sinteronto a catalyst layer. Additionally, such devices produce hydrogen usingonly the WGS reaction.

Another type of membrane is the so-called dense protonic ceramicmembrane for hydrogen separation and purification. It is based on theuse of single-phase and mixed-phase perovskite-type oxidic protonicceramic membranes for separating or decomposing hydrogen containinggases or other compounds to yield a higher value product. However, thesemembranes suffer many of the shortcomings noted previously.

Thus, it is highly desirable to develop membrane reactors that arecapable of converting liquid hydrocarbons into high purity hydrogen in asingle step.

SUMMARY OF THE INVENTION

The present invention is directed to methods and catalytic structuresthat address or reduce the weaknesses of various known catalyticstructures and which provide for a hydrogen source without carbonmonoxide. Embodiments of the present invention include methods for thesynthesis of catalytic mixed metal oxides obtained by means of thermaldecomposition of layered double hydroxides (“LDHs”) and the behavior ofthese materials on liquid hydrocarbon conversion into hydrogen using anautothermal reforming reaction.

In one embodiment, the invention provides a method for formulating anautothermal reforming (ATR) catalytic structure. The method can includethe steps of obtaining a basic solution having a basic pH, obtaining anacidic solution having an acidic pH, mixing the basic solution with theacidic solution to create a sol-gel having layered double hydroxides(LDH) precursors through co-precipitation of the metal cations with thebasic solution, heating the sol-gel at a decomposition temperature above500° C. and below a sintering temperature that would result in catalystsintering for a period of time such that the LDH precursors are at leastpartially decomposed to form a calcined material, and conducting a metalreducing step (i.e., the metal oxide is reduced to its elemental state)on the calcined material to collapse the layered crystal structureswithin the LDHs using a mixture of hydrogen on nitrogen at a temperaturein the range of about 450° C. to about 700° C. to form the ATR catalyticstructure. In one embodiment, the porosity of the ATR catalyticstructure is about 65-70%. In one embodiment, the basic solutionincludes an alkaline metal hydroxide, an alkaline metal carbonate, andwater. In another embodiment, the acidic solution includes salts andwater, wherein the salts are comprised of cations and anions, whereinthe cations are comprised of magnesium, nickel, and aluminum. In oneembodiment, the ATR catalytic structure has nanosized mixed oxideparticles dispersed throughout, wherein the nanosized mixed metal oxideshave diameters in the range of 40-300 nm. In one embodiment, the LDHshave layered crystal structures. In one embodiment, the ATR catalyticstructure can be suspended in a sol-gel and added to a tubular supporthaving a porous structure, such that the resulting apparatus is operableto convert liquid hydrocarbons into hydrogen via an autothermalreforming reaction.

In one embodiment, the nanosized mixed metal oxides have a surface areabetween 100 to 300 m²/g. In one embodiment, the nanosized mixed metaloxides are homogeneously distributed throughout the ATR catalyticstructure. In another embodiment, the nanosized mixed metal oxides arethermally stable.

In one embodiment, the basic solution has a pH value between 10 and 12.The basic solution can include a mixture of NaOH with Na₂CO₃, such thatthe basic solution has a pH value of about 12. In another embodiment,the metal cations have an absence of precious metals. Exemplary,non-limiting, precious metals include platinum, rhodium, gold, iridium,osmium, palladium, rhenium, ruthenium, and platinum. In one embodiment,the acidic solution has a total cationic concentration of approximately1.5 M. In another embodiment, the metal cations can have an aluminumconcentration of about 20 to 35 mol %.

In one embodiment, the step of conducting a metal reducing step on theLDHs is conducted at a temperature within a range from about 450° C. to700° C. In one embodiment, the decomposition temperature is within arange from about 500° C. and 600° C. In one embodiment, the ATRcatalytic structure is operable to produce a hydrogen product streamfrom a feed stream having liquid hydrocarbons through the use of an ATRreaction when incorporated into a porous tubular support.

In another embodiment, the invention provides for a method offormulating an autothermal reforming (ATR) catalyst according to thefollowing steps:

-   -   a. preparing a basic solution having a pH at or above about 10,        the basic solution including a hydroxide, preferably an alkaline        metal hydroxide, sodium carbonate and water;    -   b. preparing an acidic solution having a pH below 7, wherein the        acidic solution is prepared by combining a mixture of salts with        water, wherein the salts are comprised of cations and anions,        wherein the cations are comprised of magnesium, nickel, and        aluminum, wherein the acidic solution has a total cationic        concentration of about 1.5 M;    -   c. mixing the acidic solution and the basic solution together        for a period of time operable to form a sol-gel;    -   d. aging the sol-gel for a predetermined period of time to form        a formed solid;    -   e. washing and filtering the formed solid with water until a        generally neutral pH is reached;    -   f. drying the formed solid for a predetermined period of time to        form a dry solid;    -   g. calcining the dry solid at a predetermined temperature for a        predetermined period of time to form a calcined material; and    -   h. subjecting the calcined material to a metal reduction step by        contacting the calcined material with a mixture of hydrogen and        nitrogen gases at a temperature in the range of about 450° C. to        about 700° C. to form the ATR catalytic structure, the ATR        catalytic structure having nanosized mixed oxide particles        dispersed throughout, wherein the nanosized mixed metal oxides        have diameters in the range of about 40 to about 300 nm.

In another embodiment, the anion can include a nitrate. In anotherembodiment, the cations can further include an absence of preciousmetals. In one embodiment, the basic solution of step (a) can beprepared by combining a predetermined amount of NaOH with Na₂CO₃ suchthat the basic solution has a pH of about 12. In another embodiment, theATR catalytic structure is operable to produce a hydrogen product streamfrom a feed stream having liquid hydrocarbons through the use of an ATRreaction when incorporated into a porous tubular support. In a furtherembodiment, the hydrogen product stream is substantially free fromcarbon monoxide.

In an additional embodiment, the invention provides for an autothermalreforming (ATR) catalytic structure useful for generating hydrogen gasfrom liquid hydrocarbons. In one embodiment, the ATR catalytic structurecan include a metal oxides dispersed homogenously throughout the supportstructure. The nanosized mixed metal oxides have metal cations thatinclude aluminum, nickel, and magnesium, wherein the nanosized mixedmetal oxides have a surface area of about 100 to 300 m²/g, and thenanosized mixed metal oxides have diameters in the range of about 40-300nm. In one embodiment, the acidic solution can have a total cationicconcentration of about 1.5 M. In another embodiment, the metal oxideswithin the solution can have an aluminum concentration of about 20 mol %to 35 mol %. In another embodiment, the ATR catalytic structure isoperable to produce a hydrogen product stream from a feed stream havingliquid hydrocarbons through the use of an ATR reaction. In a furtherembodiment, the hydrogen product stream can be substantially free fromcarbon monoxide. In an additional embodiment, the ATR catalyticstructure can further include an absence of noble metal ions.

BRIEF DESCRIPTION OF THE DRAWINGS

These and other features, aspects, and advantages of the presentinvention will become better understood with regard to the followingdescription, claims, and accompanying drawings. It is to be noted,however, that the drawings illustrate only several embodiments of theinvention and are therefore not to be considered limiting of theinvention's scope as it can admit to other equally effectiveembodiments.

FIGS. 1-3 include graphical representations of experimental datacollected in accordance with embodiments of the present invention.

FIG. 4 is an embodiment of the present invention.

FIG. 5 is cross sectional view of an embodiment of the presentinvention.

FIG. 6 is an axial view of an embodiment of the present invention.

FIGS. 7-27 include graphical representations of experimental datacollected in accordance with embodiments of the present invention.

DETAILED DESCRIPTION

While the invention will be described in connection with severalembodiments, it will be understood that it is not intended to limit theinvention to those embodiments. On the contrary, it is intended to coverall the alternatives, modifications and equivalence as may be includedwithin the spirit and scope of the invention defined by the appendedclaims.

Nano sized catalysts provide advantages over their full sizecounterparts, particularly in autothermal and WGS reactions, in thatthey increase reaction activity, improve selectivity, improve reactivitytowards hydrogen production, and minimize the undesirable methanationreaction.

Carbon formation (as coke), which is a primary disadvantage of areforming process, is a kinetic issue. As such, coke production dependson the relative reaction rates of possible carbon species reactionalternatives. The reaction mechanism is largely dependent on thehydrocarbon type, operating conditions, and catalyst characteristics.For example, catalyst characteristics can influence the reactionmechanism by increasing water adsorption-dissociation rate on thecatalyst and gasification rate with respect to the C—C scission.

Catalyst characteristics are generally determined by theirphysical-chemistry, composition, structural, and textural properties,such as: active area, metal particle size, metal dispersion, andreducibility. These properties depend on metal-support interaction, andthey could be established on different stages of catalyst synthesis. Forexample, varying the material composition of a precursor, thepreparation method and/or the heat treatments (calcination orreduction), can provide desired characteristics of the catalyst foroptimum performance.

The present invention provides novel mixed metal oxides that areobtained by thermal decomposition of layered double hydroxides (LDHs)and offers the opportunity to control a catalyst's active site natureand its environment, as well as catalyst texture and stability. LDHs area unique class of layered materials having positively charged layers andcharge balancing anions located in the interlayer region. Typically,LDHs can be synthesized by the co-precipitation of metallic salts with aconcentrated alkaline solution. An alternative method for thepreparation of LDHs is through the sol gel method. Thermal treatments ofthe mixed metal oxides resulting from the LDHs prepared by the sol gelmethod can lead to materials that demonstrate synergetic effects betweenthe elements in the mixed metal oxide structure, and after appropriateactivation treatment, give rise to well dispersed metal particles likesupported metal catalysts, with the possibility of controllingmetal-support interaction during the synthesis stage.

In one embodiment of the present invention, a new catalytic structurethat can be used for reforming reactions is provided. In one inventivemethod of preparing said catalytic structure, several different types ofLDHs can be prepared using a sol-gel precipitation method. In oneembodiment, two aqueous solutions are prepared; one acidic and onebasic. In certain embodiments, the sol is base catalyzed. The acidicsolution preferably contains one or metal salts, such as magnesiumsalts, and specifically nitrates and salts of the same for nickel andaluminum using a desired Al/(Ni+Mg+Al) ratio. The pH of the acidicsolution is preferably within the range of about 4 to about 6. In oneembodiment, the total cationic concentration of the metal salts in theacidic solution is about 1.5 M. The basic solution can be obtained bymixing suitable amounts of sodium hydroxide and sodium carbonate inorder to maintain a ratio of carbonate ions to nickel, magnesium andaluminum ions of around 0.7 and a pH for synthesis of approximately 12.

After the solutions are prepared, they can then be added into a largemixing device and put into mechanical agitation for an appropriateperiod of time to form a sol-gel. In one embodiment, the appropriateamount of time is determined by measuring the concentration of metaloxide. In one embodiment, 95% concentration of metal oxide isacceptable. At concentrations at or above 95%, a preferred amount ofreactant has been consumed, and agitation can be stopped. In oneembodiment, ICP-AES analysis can be used to verify the metal oxideconcentration. In one embodiment, this mixing step can be for up toabout 5 hours. Following the mixing step, the obtained sol-gel is leftto age at a temperature of between about 50 and 75° C., alternatively ata temperature of about 60° C. for an appropriate amount of time, whichcan be up to about 6 hours, alternatively up to about 10 hours,alternatively up to about 15 hours, or greater than 15 hours. Duringaging, at least a portion of liquid within the sol-gel evaporates. Inone embodiment, rapid aging is not preferred as it can detrimentallyimpact gel characteristics. The resulting solid can then be filtered andwashed with distillated water until the wash water has a neutral pH ofabout 7. Water washing is desirable to remove all undesired andconverted reactants. Those of ordinary skill in the art will understandneutral pH to including a pH within the range of about 6.5 to 7.5. Next,the resulting washed solid LDH material can be dried at about 100° C.for an appropriate amount of time, which can be up to about 5 hours,alternatively up to about 10 hours, alternatively up to about 15 hours.In one embodiment, the LDH material can then be calcined in air attemperatures up to 550° C., alternatively at a temperature of betweenabout 400° C. and 600° C., alternatively at a temperature between about400° C. and 500° C., alternatively at a temperature between about 500°C. and 600° C., for a period of time, which can be up to about 5 hours,alternatively up to about 8 hours, alternatively up to about 12 hours.The resulting solid product is a catalyst precursor of a mixed metaloxide that includes Ni/Mg/Al. Supported metal catalysts suitable forautothermal reforming reaction of liquid hydrocarbons into hydrogen canbe obtained from these precursors after a metal reducing step bycontacting with a gas mixture that includes hydrogen and can alsoinclude nitrogen at a temperature of about 500° C., alternativelybetween about 400° C. and 550° C., for a period of time, which is up toabout 5 hours. In one embodiment, the gas mixture contains about 5%molar hydrogen, alternatively between about 1 and 10% molar hydrogen,alternatively between about 5 and 10%, alternatively between about 5 and15%, alternatively between about 10 and 20% molar hydrogen.

During the metal oxide reduction step, at temperatures greater thanabout 300° C., the LDL structure will collapse, and in the presence ofhydrogen, the metal oxide can be reduced to its elemental state. Thecollapsed structure was confirmed using XRD (and is shown in FIGS. 1-3).In general, the collapsed structure has an increased density, relativeto the starting material, and a porosity that is sufficient fordiffusion of reactants into the catalyst to undergo the desiredreactions. In certain embodiments, the porosity of the collapsedmaterial can remain consistent with the starting material, having aporosity that is in the range of about 65-70%, which in turn canfacilitate the desired reforming reactions.

Experimental Design for Catalytic Structure:

The following synthesis parameters were investigated to obtain catalyststructures in accordance with embodiments of the present invention:

-   -   Al/Mg mole ratio between 0.2-5 in the form of mixed metal oxides    -   Calcination temperature: 500° C.-600° C.    -   Reduction temperature: 500° C.    -   Concentration of Hydrogen: 5 to 20% by mole    -   Nickel incorporation method: Co-precipitation simultaneously        with magnesium and aluminum or impregnation on Al—Mg support.

In order to study the synthesis parameter effects on catalyticcharacteristics of these materials, different series of samples wereprepared. In each different sample, only one parameter was changed. Thebaseline material, identified as CP10C, had the following parameters:

-   -   Nickel content: 10 wt %    -   Cation ratio [Al/(Al+Mg+Ni)]: 0.20    -   Calcination temperature: 550° C.    -   Reduction temperature: 500° C.    -   Nickel incorporation method: co-precipitation with magnesium and        aluminum.

To determine the Al/Mg ratio effect on catalytic behavior, severalcatalysts were prepared wherein the Al/Mg ratio was varied between 0.2and about 5.

To determine the calcination temperature effect on catalytic behavior,several mixed oxides catalysts were prepared wherein the calcinationtemperatures were varied between about 400° C. and about 800° C.

To determine reduction temperature effect on catalytic behavior, severalcatalysts were prepared wherein the reduction temperatures was variedbetween about 450° C. and about 700° C.

To determine the effect of the nickel incorporation method on catalyticbehavior, several catalysts were prepared by co-precipitation and byimpregnation. Catalysts marked with “LO” were prepared by impregnation,and catalysts marked with a “CP” were prepared by co-precipitation.

The results and characteristics for each of the materials were measuredusing several different techniques in order to correlate catalyticactivity with corresponding structural properties. Catalyst supportsurface area was measured using the BET technique employing nitrogenphysisorption at the temperature of liquid nitrogen in a QuantachromeAutosorb-1C instrument. The percentage of metal loading was measured byInductively-Coupled Plasma Atomic Emission Spectrometry (ICPAES). X-rayDiffraction (XRD) was carried out using Rigaku Miniflex diffractometeremploying a Cu-Kα radiation source (30 KV/15 mA). The average particlesize for different phases present on each state of the sample (calcined,reduced or used) was estimated by the Scherre equation, which isreproduced below:

$\tau = \frac{K\;\lambda}{\beta cos\vartheta}$where K is the shape factor, λ is the x-ray wavelength, β is the linebroadening at half the maximum intensity (i.e., full width at halfmaximum) in radians, θ is the Bragg angle; and τ is the mean size of theordered (crystalline) domains.

Nickel surface area can be measured using well established advancedanalytical techniques, such as hydrogen chemisorptions complemented bytemperature programmed reduction (TPR) where it is possible to estimatetotal metal surface area, size of metal particle and metal dispersion.

Temperature programmed reduction (TPR) was carried out using a 5% H₂/N₂mixed gas flowed at 50 mL/min and used to study the reducibility of thecalcined samples by means of hydrogen consumption. From TPR curves, thetemperature at which a maximum curve appears (nickel reductiontemperature) can be used to determine the most efficient reductionconditions. The degree of reduction can be estimated by comparing TPRcurves corresponding to calcined and reduced samples. Coke formation onsamples during activity test can be determined by analysis of thetemperature-programmed oxidation (TPO). Autothermal reforming reactionstudy was carried out in a down low fixed-bed reactor catalytic system.

Approximately 1 gram of the catalyst was used having particles the sizeof about 30-50 mesh without dilution. The catalyst tested waspre-reduced in an H₂ gas stream (5%) at a temperature of about 500° C.for about 5 hours. Two thermocouples controlled the operationtemperature, wherein a first thermocouple was positioned in the oven andthe other thermocouple was positioned at the center of the catalyst bed.All reactants were introduced from the top of the reactor. Air was usedas a source of oxygen. Two HPLC pumps fed water and hydrocarbon, whichwere mixed and evaporated before being introduced into the reactionregion. The pressure of the reaction was maintained by a backpressureregulator connected with a precision gauge to read the pressures. Timezero was measured at the point when the reactor contents were heated tothe chosen reaction temperature, which usually took about 45 minutes.Gas samples were removed through the gas sampling system throughout thelength of the test once the reaction temperature was reached. Effluentfrom the reactor was cooled in a double pipe condenser to condense thecondensable vapors. The gas and liquid samples were analyzed by gaschromatography.

The autothermal reforming reaction was investigated at differentreaction conditions as follows:

-   -   Hydrocarbon feed: n-octane,    -   Reaction Temperature: 500° C., 550° C. and 600° C.    -   Total Pressure: 3 bars    -   Weight Hourly Space Velocity (WHSV): 5000 h⁻¹    -   Oxygen/Carbon (molar ratio): 0.5    -   Steam/Carbon (molar ratio): 3

The general reaction procedure included packing the catalysts into thereactor, starting the reactor temperature controller and then, followingcatalyst reduction, setting the desired temperature condition.Approximately, fifteen minutes before the test began, steam was passedthrough the catalyst bed. Following this, the air flow meter with thewater and hydrocarbon pumps were turned on all together at theestablished conditions.

Each experiment ran for approximately eight hours. The first two hourswere conducted at temperature of about 500° C., and then increased to atemperature of about 600° C. for a further two hours, and againincreased to a temperature of about 700° C. for two hours and finallyreturning to a temperature of about 500° C. for the final two hours.Every thirty minutes, a reformate gas sample was taken for analysis todetermine the molar concentrations of H₂, O₂, CH₄, CO and CO₂.Volumetric flow (ml/min) and density (g/ml) of the liquid reformateproducts were also tested, and were passed through a GasChromatography-Flame Ionization Detector (GC-FID) assay.

Two Shimadzu GC-17A Gas Chromatography (GC) units were used to estimatethe composition of the product gases and liquid reformate collected fromthe ATR reaction. The GC units were equipped with thermal conductivity(TCD) and flame ionization (FID) detectors. Specifically, the GC-TCD wasused to evaluate the presence of H₂, O₂, N₂, CH₄, CO and CO₂ using a ⅛inch 6 ft Carbosphere 80/100 packed column. The light and heavy liquidhydrocarbons were analyzed using a GC-FID coupled with a 100 m×0.25 mmID BPI-PONA capillary column. Liquid product identification was alsocarried out in a Shimadzu GCMS-QP5050A mass spectrometer equipped with aDB-5 column.

For the GC-TCD method, the injector was held at 100° C. and detectortemperature was 150° C. The oven temperature was initially maintained ata temperature of 40° C. for 5 minutes, and then increased to atemperature of 120° C. at a rate of 5° C./min. In the case of theGC-FID, the injector was held at 280° C. and detector temperature was320° C. The oven temperature was initially maintained at a temperatureof 60° C. for 5 minutes, and then increased to 280° C. at a rate of 10°C./min.

Referring to FIG. 1, x-ray diffraction spectra for 4 samples prior tocalcination are provided. Samples LO1 and LO2 were prepared byimpregnation, and samples CP1 and CP2 were prepared by co-precipitation.The samples had compositions of about 20.5% by weight NiO, 5.5% byweight MgO, and 75% by weight Al₂O₃. The samples were all air-dried at110° C. for approximately 5 to 15 hours. As shown in the x-raydiffraction spectra stacked plot shown in FIG. 1, the individual spectrafor each sample are relatively similar, thus demonstrating that thesynthetic methods utilizing impregnation and co-precipitation, whetherat a pH of 10 or 12, produces material having like crystallographicstructure. Additionally, FIG. 1 also displays the LDH structure ofdifferent aluminum/magnesium crystallite sizes, along with thecrystallinity at different catalyst synthesis conditions.

FIG. 3 shows x-ray diffraction spectra for four calcined catalystsamples having a composition of about 20.5% by weight NiO, about 5.5% byweight MgO, and about 75% by weight Al₂O₃, wherein the samples werecalcined at a temperature of about 650° C. for a period of about 8hours. The materials correspond to the calcined versions of thematerials shown in FIG. 1. The spectra show that samples LO2C and CP2Cboth have peaks at about 38, 46 and 66, which correspond to mixed metaloxide (as opposed to LDH structures). Sample CP2C also includes a slightpeak at about 60, corresponding to metal oxide. In contrast, sample LO1Chas peaks at about 38, 42, 46, 63 and 66, and sample CP1C has peaks atabout 38, 42, 62, and 66, although the peak at 66 is of a lowerintensity.

Based on the experimental results, which are shown in FIGS. 1-3,Applicants surprisingly discovered the following:

(1) Catalysts obtained by calcination/reduction of LDH materials provedto be suitable catalyst materials for the autothermal reforming ofliquid hydrocarbons for hydrogen production. These catalyst materialsexhibit highly dispersed metallic crystallites or grains that are stableinside a matrix, and advantageously have a high surface area.

(2) The catalyst materials obtained in accordance with variousembodiments of the present invention can be used in catalytic membranesto convert liquid hydrocarbons, steam, and air into hydrogen and carbondioxide. The catalytic activities test produced nearly total liquidhydrocarbon conversion with large levels of hydrogen produced withoutany carbon monoxide production, thus indicating an excellent balance ofATR and WGS combined reactions. Additionally, the WGS reaction wasfavored over the methanation reaction, resulting in a reduced amount ofmethane gas being produced. Moreover, coke formation was also reduced,which improves overall catalytic activity, thus leading to longer runtimes.

(3) Nickel incorporated into the catalyst materials by co-precipitationresults in more highly and finely dispersed nickel particles than theimpregnation method.

(4) Synthesis parameters have an important influence on metal-supportinteraction. When stronger metal-support interactions are favorable, thecatalyst structure is improved. Additionally, catalyst performance isstrongly linked to the nickel particle size, as well as the selectedsupport material and its properties.

As noted previously, prior art catalytic membranes were used to produceand subsequently remove hydrogen through hydrogen permeable metalmembranes, such as palladium or palladium alloys. These membranes, inparticular palladium or palladium alloys, however, are expensive,sensitive to sulfur compounds, and difficult to sinter with or co-sinteronto a catalyst layer. Additionally, such devices typically producedhydrogen only by the WGS reaction.

Some embodiments of the current invention integrate catalysts useful forboth the ATR reaction and the WGS reaction, as well as providing ahydrogen permeable membrane that does not include palladium or apalladium alloys, which reduces costs while increasing the overall H₂yield of the process. In one embodiment, the ATR structure describedherein can be incorporated into a catalytic membrane reactor (CMR)assembly, which is operable to perform the ATR reaction, the WGSreaction, and remove hydrogen.

In one embodiment of the invention, a method for the production ofhydrogen by oil reforming processes reduces the overall cost ofproducing high quality hydrogen from liquid oil, as compared to theprior art processes. In one embodiment, the method can include agasification, steam reforming, partial oxidation, autothermal reforming,or like step, depending on the characteristic of the oil feed processedto obtain synthesis gas. This synthesis gas requires subsequent cleaningas it is shifted to produce additional hydrogen gas. In one embodiment,one or more of these steps can be combined to improve efficiency. Incertain embodiments of the present invention, additional noveltechnologies, such as membrane separation and catalytic reactors, havebeen developed that can help to solve these needs. Additionally, someembodiments of the present invention are also operable to removeunwanted products, such as CO₂, to thermally provide one stream that canbe used in a secondary process, such as in enhanced oil recovery fromdepleted oil reservoirs.

In one embodiment, methods for producing hydrogen from alternatehydrocarbon sources of oil (such as, gasoline, kerosene, diesel,petroleum coke, heavy residues, and the like) involves the step of firstreacting oil with oxygen and/or steam to produce a gas mixture thatincludes carbon monoxide, carbon dioxide and hydrogen. The carbonmonoxide can then react with steam to produce additional hydrogen andCO₂. Finally, the hydrogen and CO₂ can be separated, either by removingthe CO₂ from the mixture or by removing hydrogen from the mixture. Theremoval of at least a portion of the hydrogen gas will advantageouslyshift the reaction equilibrium toward the product side, allowing alowering of reaction temperature and use of a decreased amount of steam.An optional cleaning step can also be employed, but CO contamination canbe controlled to make such contamination negligible. For example,concentrations of CO as low as 0.001% can be obtained, making thehydrogen substantially free of carbon monoxide. In certain embodiments,all of the hydrogen produced can be removed.

The combination of reaction and separation processes, as describedherein for the CMR assembly built in accordance with embodiments of thepresent invention, offers higher conversion of the reforming reaction atlower temperatures due to the step of removing hydrogen gas from thesteam reforming and WGS equilibrium reactions. Thus depending on thefeed oil composition used, a membrane reactor as part of an engineeringprocess, can allow one step reforming and/or partial oxidation with WGSreaction and parallel hydrogen separation. Unlike conventional prior artprocesses, the CMR assembly equipped with the catalysts describedherein, benefits from high pressure operation due to the increasedhydrogen partial pressure differences across the membrane, which act asthe driving force for hydrogen permeation. For example, the pressure onthe permeate side can be atmospheric or under vacuum. When a higherpressure on the retentate side is applied, the pressure differenceacross the membrane acts as a driving force for hydrogen to permeatethrough the membrane. The higher the pressure difference, the higher theamount of hydrogen permeating through the membrane. Those of ordinaryskill in the art will recognize that the mechanical property of themembrane used will create a practical limit as to the pressure that canbe applied on the retentate side.

Another embodiment of the present invention discloses a method ofmanufacturing a catalytic coated silica membrane for the conversion ofliquid petroleum hydrocarbon fuels into high purity hydrogen. Theembodiment includes the step of providing a membrane tube, whichincludes an outer surface covered with an active silica layer that ishighly permeable to hydrogen. A mixed metal oxide catalyst, which caninclude at least one metal selected from the group consisting ofrhodium, platinum, nickel, ruthenium, palladium, rhenium, iridium, andcombinations thereof, can be deposited within the pores of the aluminaframework of the membrane tube. During use, air, steam, and liquidhydrocarbons are transported through the membrane and are in intimatecontact with the metal sites. Following activation of the catalystmembrane, the feed components react and form hydrogen through acombination of ATR and WGS reactions. High purity hydrogen can beproduced by separation of the hydrogen gas from the product mixturethrough the hydrogen permeable membrane deposited on the CMR assemblyouter surface. Certain embodiments of the present invention bring anumber of benefits to hydrogen production from liquid petroleumhydrocarbon fuels. These improvements can include high conversion of theliquid hydrocarbons, high molar yield of the hydrogen produced, lowmolar yield of the residual methane and low catalyst deactivation as aresult of the shifting of reaction equilibrium to favor the forwardreaction to nearly 100% at a much lower operating temperature rangebetween about 500° C.-550° C. Typical hydrogen purity in this one-stepconversion of liquid petroleum hydrocarbon to hydrogen can range fromabout 96 to 99% molar concentration, alternatively at least about 97%molar concentration, alternatively at least about 98% molarconcentration, or alternatively between about 97 and 99% molarconcentration.

In one embodiment, the CMR assembly in accordance with an embodiment ofthe present invention can be selectively permeable to hydrogen andproduce a hydrogen-rich permeate product stream on the permeate side ofthe membrane and a carbon dioxide rich product retentate. The CMRassembly can be used to produce a hydrogen-rich permeate product streamthat is greater than about 99% by volume hydrogen. The CMR assembly canbe a composite ceramic material having an outer hydrogen transport andseparation layer, which in one embodiment is a metal doped silica. TheCMR assembly can be also composed of one or more inner catalytic layers.The ATR and WGS reactions occur on the inner metal layers of the CMRassembly and the produced hydrogen can be transported and removedthrough the outer metal-doped silica layer.

In one embodiment, the metal catalysts are capable of catalyzing theconversion of hydrocarbon fuels to hydrogen and carbon oxides (CO andCO₂). In one embodiment, the catalyst can include one or more of thefollowing metals: nickel, ruthenium, platinum, palladium, rhodium,rhenium, and/or iridium.

The CMR assembly can include a stainless steel vessel that surrounds thetubular support and is positioned between a pair of high temperaturematerial shells. In one embodiment, the high temperature material shellscan help to provide sealing, manifolding, expansion support, separatedregions for the catalytic reactions, delivery of pressurized feedstock,support of the membrane, and removal of product gases.

In accordance with one embodiment of the present invention, thestainless steel vessel was constructed of 316 grade stainless steel. Inan exemplary embodiment, a cylindrical α-alumina tube having an outerdiameter of about 10 mm and an inner diameter of about 8 mm with anaverage pore size of 1.3 micro meters and an approximate porosity of0.55 was used as ceramic tube support for the membrane. For thepreparation of the active catalytic sites and the membrane separationpores, the following procedure was adopted:

First, the α-alumina ceramic tube was soaked in a boehmite sol-gel,having a concentration in the range of about 5 to 10% by weightboehmite. After removing the excess boehmite by drip drying, the ceramictube was dried by natural convection at room temperature for severalhours and then at about 110° C. for up to about 10 hours. It wasfollowed with calcination at a temperature of about 500° C. thatresulted in the bimodal distribution for the original α-aluminastructure into alpha and gamma phases. During the calcination step, thenitrates and other gaseous oxides were evaporated and converted to themetal oxide form of the active metals and support structure (e.g., NiO,MgO, and Al₂O₃). For the materials used, the preferred calcinationtemperature was between 500° C. and 800° C. At temperatures above 800°C., it was discovered that the metal oxide structure changed to a dense,low surface area material, which greatly reduced its catalytic activity.To introduce the catalyst metal active sites to the tubular structure,the ceramic tube was soaked in a solution containing salts of metalspecies, such as nickel, ruthenium, platinum, palladium, rhodium,rhenium, iridium, and combinations thereof, at the appropriateconcentration. Drying and calcination steps were subsequently performed,resulting in a uniform metal distribution within the core of the ceramictubular structure. The final step was to dope the outer surface of themetal loaded catalytic ceramic tube with a silica colloidal sol-gel toproduce an even thickness and uniform pore size hydrogen permeablelayer. In one embodiment, this process can be repeated several times toremove large pores that might result in pinholes in the final membrane.In one embodiment, the silica colloidal sols were coated on the ceramictube by a hot coating method in which the tube was first heated to atemperature of up to around 180° C. and then subsequently and quicklycontacted with a wet cloth containing the sols. This hot coating methodhelped make the sol dry more quickly, preventing the sol frompenetrating deep into the pores. The catalytic membrane can be cut inpieces and then characterized by Scanning Electron Microscopy (SEM),nitrogen adsorption, porosimetry and hydrogen adsorption.

In order to test the hydrothermal stability of the silica membrane, astability experiment was carried out at about 500° C. using severaldifferent molar ratios of steam and liquid hydrocarbon. Thepermeability-selectivity of the ceramic tube to hydrogen was measured attimed intervals. In general, a very slow drop in the permeationcharacteristics due to humidity was observed. Typically, at hightemperature, silica interacts with steam causing densification of thematerial. Within the temperature ranges used in the experiment (e.g.,between about 500° C. and about 600° C.), the membrane did not show anysubstantial change in its permeation properties. In addition, theinfluence of the molar ratio of water (steam) to hydrocarbon on liquidhydrocarbon conversion within the temperature range of 500° C. to 550°C. showed that the conversion was enhanced at high steam to hydrocarbonmolar ratios, such as a steam to hydrocarbon molar ratio of greater thanabout 3:1, or alternatively between about 2:1 and about 3:1. Thehydrocarbon feed reacted very effectively with steam and air through theATR reaction steps. A fraction of the CO produced during the ATRreaction subsequently reacted with steam in the WGS reaction, leading toincreased hydrogen yield. In-situ removal of hydrogen by the membranefurther enhanced conversions of both of these reactions. In addition,enhancement in the rate of CO removal by the WGS reaction can reducecoke deposition caused by the so-called Boudouard reaction.

In FIG. 4, CMR assembly 10 includes: stainless steel vessel 12, shells14, ceramic tube 20, feed stream inlet 30, retentate output 40, sweepstream inlet 50, and permeate output 60. The feed stream (not shown),which includes liquid hydrocarbons, steam, and the oxygen source, entersCMR assembly 10 thru feed stream inlet 30. The feed stream then movesinto ceramic tube 20 where it undergoes both ATR and WGS reactions,thereby forming hydrogen and carbon dioxide. The hydrogen that permeatesthru the permeable membrane of ceramic tube 20 is removed by a sweepgas, such as nitrogen, argon, steam, and combinations thereof, thatenters thru sweep stream input 50 and exits permeate output 60.

FIG. 5 shows an embodiment in which ceramic tube 20 has ATR layer 22 andWGS layer 24 disposed within the tube in two separate layers. Those ofordinary skill in the art will recognize that the ATR catalysts and WGScatalysts can also be disposed within the same layer. In the embodimentshown in FIG. 2, the feed stream enters ceramic tube 20 where it reactswith ATR catalysts disposed in ATR layer 22. This reaction converts thehydrocarbon, oxygen, and steam in the feed stream into hydrogen gas andcarbon monoxide gas. The carbon monoxide can react with additional steamand the WGS catalysts disposed in WGS layer 24 to form carbon dioxideand additional hydrogen gas. The hydrogen gas then travels radiallyoutward through the pores of γ-alumina layer 26 and then throughmembrane 28. The carbon dioxide gas cannot permeate membrane 28, andexits ceramic tube 20 on the retentate side of membrane 28. Unconvertedhydrocarbons and carbon dioxide exit ceramic tube 20 through retentateoutput 40 of FIG. 4.

FIG. 6 is an axial view of ceramic tube 20 as shown in FIG. 5 along line6-6.

Experimental Design for CMR Assembly:

Commercial alumina tubes (obtained from Noritake, Japan) of 100 mmlength and 10 mm ID were used for the CMR assembly. The tubes includeα-alumina having a porosity of about 30% and average pre size of about0.5-1 μm. A γ-alumina layer having a pore size of about 4 nm wasdisposed on the outer surface of the tubes. A metal doped sol-gel wasprepared by mixing 120 g of tetraethyl orthosilicate (TEOS) in 600 gethanol to form a solution. An acid solution was then synthesized bydissolving 14.84 g cobalt nitrate hexahydrate (Co(No₃)₂.6H₂O) in 51.77 gof 30% aqueous H₂O₂. Subsequently, both solutions were vigorously mixedtogether by stirring for about 3 hours in ice-cooled bath. The tubeswere then externally coated with a stable Si—Co—O solution using acontrolled immersion time of 1 minute and withdrawn speed of 2 cm/min.Sintering was then carried out at about 600° C. for about 4 hours at aheating rate and a cooling rate of 0.7° C./min.

The internal area of the alumina tube was wetted with a sol-gel thatincludes a boehmite solution having a concentration in the range of 5-10wt % of aluminum. Excess solution was wiped off, and the tube was driedfor several hours at a temperature of about 110° C., and then fired at atemperature of about 500° C. for about 8 hours, which resulted in aninlet tube wall of γ- and α-alumina having a bimodal distribution.Subsequently, this internal layer of the tube was soaked withchloroplatinic acid and rhodium chloride solutions (0.5 wt % platinumand 2.5 wt % rhodium, respectively), followed with drying, calcination,and reduction under hydrogen to finely disperse the active metalparticles over the γ- and α-alumina pores.

In one embodiment for the preparation of the chloroplatinic acid andrhodium chloride solutions, the following procedure was used: TEOS wasmixed together with the metal solutions, water and ethanol according tothe following:

-   -   TEOS: 0.9 parts    -   Pt—Rh: 0.5 to 2.5 parts    -   Water: 4 parts    -   HNO₃: 0.01 parts    -   Ethanol: 10 parts        Afterwards, the resulting solution was hydrolyzed for about 12        hours at room temperature. After adding an appropriate amount of        a 0.1 M solution of nitric acid, the solution was left at room        temperature for about 5 hours to convert into a colloidal sol        solution. It was then used to coat the internal wall of the        membrane tube.

A silica metal doped hydrogen separation layer having a thickness ofabout 1 μm was formed on the external wall of the membrane tube. Aninternal, highly dispersed Rh—Pt catalytic layer, having a metaldispersion of about 65%, was impregnated on the bimodal layer (whichincludes α- and γ-alumina) to act as a catalytic layer for thesimultaneous ATR and WGS reactions to occur in the internal wall of themembrane tube.

Experimental Results of CMR Assembly:

A set of initial experiments were conducted using the CMR assemblydescribed above and shown in FIGS. 4-6, and a commercially availableautothermal catalyst (FCR-71D autothermal catalyst provided by SudChemie). One gram of the commercially available catalyst was used undersimilar experimental conditions tested on a fixed bed reactor (FBR)during catalyst selection: Temperature of 500° C.; Pressure of 3 bars;oxygen:carbon ratio of 1:2; steam:carbon ratio of 3:1; and a hydrocarbonfeed of commercial gasoline. FIG. 7 shows the total gas productdistribution at the end of both the retentate and the permeate side ofthe membrane. Hydrogen concentration on the permeate side was observedto be as high as 92% by volume. A 77 wt % recovery of hydrogen wasobtained, with the remainder of the product stream being a mixture ofcarbon monoxide, carbon dioxide, methane and unconverted oxygen. A totalconversion of about 98% of commercial gasoline was achieved.

The hydrogen yield produced using a fixed bed reactor system and whenusing a CMR assembly was compared and can be seen in FIG. 8. As shown inFIG. 8, embodiments of the present invention utilizing the CRM assemblyprovide a distinct advantage over fixed bed reactors due to shifting theconditions of the reaction equilibrium.

In FIGS. 9 and 10, the temperature was increased to 550° C. Notably, theCMR reactor had a higher selectivity of hydrogen (about 15% more) ascompared to the fixed bed operation. Additionally, as shown by comparingFIG. 9 and FIG. 10, the CMR operation reduced the production of methaneby approximately 40%, as compared with the fixed bed reactor.

FIG. 11 shows comparisons of the average concentrations of hydrogen,carbon monoxide and methane in the product streams for fixed bedoperation and CMR operation as function of temperature (500° C. and 550°C.). Comparing both fixed bed reactor and CMR operations, an increasedamount of hydrogen is produced at an operating temperature of 500° C.than is produced at an operating temperature of 550° C. Additionally, atincreased temperatures, the reduction in methane produced is verysignificant. On the retentate side of the reaction, analysis of the COdemonstrates nearly the same values for both operations.

The strength of the silica membrane and the hydrogen separation throughthe retentate was tested for both the CMR and FBR reactors at 600° C.,using the same reaction parameters as provided above with respect toFIGS. 7-9. The experiment was conducted for about 6.5 hours, however,after completion, leaks were observed through the membrane and throughthe graphite sealing. FIG. 12 represents a comparison of the results onhydrogen selectivity are compared as function of temperature and type ofreactor used. As shown in FIG. 12, hydrogen is produced in amountsgreater than the thermodynamic equilibrium conditions at temperatures of500° C. and 550° C. Additionally, both CMR and fixed bed reactors showdecreased hydrogen selectivity with increased temperature.

FIGS. 13 and 14 compare the selectivity of two reactors to methane andcarbon monoxide. When using CMR operation, the methane yield issignificantly reduced compared with the values for FBR. This behaviorcan be attributed to the fact that in using the CMR device, the thermalcracking reaction of the fuel and the Boudouard reaction that causesmethanation, are inhibited because hydrogen is continually being removedthrough the silica membrane, which in turn causes more carbon monoxideto be consumed. On the other hand, the carbon monoxide yield between thetwo operation modes remains practically unchanged. As such, CMRoperations increase yields of hydrogen, reduce yields of methane, andhave a negligible effect on carbon monoxide production.

Table I presents the comparison of gasoline conversion, total hydrogenproduced and the percentage of hydrogen recovered through the membraneas function of the reaction temperature. Increased reaction temperaturefavors the hydrogen recovery from the membrane, although the totalhydrogen yield decreased.

TABLE I Composition Production for CMR as a Function of TemperatureTemperature Gasoline Hydrogen % Hydrogen (Celsius) Conversion (%)produced (mol %) Recovered 500 89 72 90 550 94 64 94

FIG. 15 shows the levels of catalyst deactivation for the two types ofreactors. As shown in FIG. 15, carbon deposition on the catalyst surfaceis lower for CMR operation as compared with FBR operations.Additionally, carbon deposition on the catalyst surface is nearlyindependent of temperature for CMR operation, contrasting with FBRoperation where carbon deposition increased by nearly 100% as thetemperature was increased from about 500° C. to about 600° C.

Coke deposit selectivity can be seen in FIG. 16, which shows thetemperature programmed oxidation (TPO) analysis of the coked catalystsusing CMR and FBR operations. The results indicate that coke depositionfrom catalysts used in CMR operation mainly covers the support sites ofthe catalysts, thereby leaving the metal active sites clean and freefrom deactivation. FBR, on the other hand, shows deposition on supportsites and also on metal sites and in the vicinity of metal-supportsites. This coke deposition can break the reaction chain to producehydrogen as reflected by the high methane and light unsaturatedhydrocarbons that were found during FBR operations.

FIGS. 17 and 18 show the gas selectivity for permeate and retentate forthe membrane sample CMR-1. In this case, hydrogen purity in the permeatewas greater than about 95 mol %. Relatively low amounts of CO andmethane were produced, which is an indication that the methanationreaction is suppressed and the WGS reaction is enhanced.

FIGS. 19 and 20 show the gas selectivity for permeate and retentate formembrane sample CMR-2. CMR-2 is a tube having similar physicalcharacteristics to that of CMR-1, and it was used to showreproducibility. Similar to CMR-1, CMR-2 displayed a high hydrogenpurity in permeate of around 90 mol %. CMR-2 also showed good membranestability and very low production of CO and methane, which is anindication that the methanation reaction is suppressed and the WGSreaction is enhanced.

FIGS. 21 and 22 show the gas selectivity for permeate and retentate formembrane sample CMR-3. CMR-3 is a tube having similar physicalcharacteristics to that of CMR-1 and CMR-2; however, it was from adifferent supplier, and was used to show reproducibility. Similar toCMR-1 and CMR-2, use of CMR-3 resulted in a high hydrogen purity inpermeate of around 90 mol %. CMR-3 also showed good membrane stabilityand very low production of CO and methane, which is an indication thatthe methanation reaction is suppressed and WGS reaction is enhanced.

FIGS. 23-25 show the hydrogen purity concentration on permeate andretentate for each of the 3 membranes; CMR-1, CMR-2 and CMR-3. FIGS.23-25, show that CMR-1 has the largest differences in hydrogen puritybetween the permeate and the retentate, indicating that CMR-1 is themembrane with the best hydrogen recovery rates (i.e., about 82%), whichis shown in FIGS. 26 and 27.

There are various alternative uses for the CMR assembly as constructed.For example the membrane system can be used as a hydrogen extractor forremoving hydrogen from refinery gas streams such as the catalyticcracker off gas, hydrodesulfurization gas, etc. Other uses can alsoinclude hydrogenation, dehydrogenation, fuel cells, and the like.

By integrating a hydrogen perm-selective silica layer with a reformingcatalytic layer, an efficient and compact CMR assembly for reformingliquid hydrocarbon fuels into hydrogen has been shown. The system showsincreased performance in the reforming of liquid hydrocarbons, atcomparatively lower operating temperatures and lower steam to carbonmolar ratios, than is typically the case for conventional fixed-bedreforming processes. The process efficiency, to a large extent, dependson various process parameters. Under optimized conditions, a nearly 30%improvement from the equilibrium conversion levels was achieved as aresult of continuous hydrogen removal from the product stream throughthe CMR assembly that employs a hydrogen permeable silica membraneintegrated with the catalyst layer. The system offers a promisingalternative to conventional hydrogen permeable membrane reactors thatare typically packed with either reforming or WGS catalysts.

Use of terms such as “first” or “second” are strictly for labelingpurposes and do not connote order of operations.

While the invention has been described in conjunction with specificembodiments thereof, it is evident that many alternatives,modifications, and variations will be apparent to those skilled in theart in light of the foregoing description. Accordingly, it is intendedto embrace all such alternatives, modifications, and variations as fallwithin the spirit and broad scope of the appended claims. The presentinvention may suitably comprise, consist or consist essentially of theelements disclosed and may be practiced in the absence of an element notdisclosed.

We claim:
 1. A method for producing a hydrogen rich permeate stream, themethod comprising the steps of: introducing a feed stream comprising aliquid hydrocarbon, steam, and oxygen to a catalytic membrane reactorassembly at an operating temperature and operating pressure, thecatalytic membrane reactor assembly comprising a catalytic structurehaving a porosity in the range of 65% to 70% and metal particleshomogeneously distributed throughout the catalytic structure, theoperating temperature and operating pressure being selected such thatthe feed stream undergoes an autothermal reforming (ATR) reaction and asubsequent water-gas-shift (WGS) reaction, wherein the metal particlesare formed by conducting a metal reducing step on the catalyticstructure having mixed metal oxide particles, the mixed metal oxideparticles formed by thermal decomposition of layered double hydroxidesin a sol-gel added to a bimodal alumina support structure of thecatalytic membrane reactor assembly; and producing a permeate stream anda retentate stream using the catalytic membrane reactor assembly,wherein the permeate stream comprises hydrogen, wherein the retentatestream comprises carbon dioxide and unreacted hydrocarbons.
 2. Themethod of claim 1, comprising removing at least a portion of thepermeate stream from the catalytic membrane reactor assembly to therebyshift the reaction equilibrium of the ATR reaction and the WGS reactiontoward the retentate, wherein removing at least a portion of thepermeate stream allows for lowering the operating temperature of thecatalytic membrane reactor assembly and decreasing the amount of steamin the feed stream.
 3. The method of claim 1, wherein the liquidhydrocarbon comprises gasoline.
 4. The method of claim 1, wherein thehydrogen is substantially free of carbon monoxide.
 5. The method ofclaim 1, wherein the operating temperature of the catalytic membranereactor assembly is within a range of 500° C. to 550° C.
 6. The methodof claim 1, wherein the catalytic membrane reactor assembly comprises ahydrogen permeable membrane.
 7. The method of claim 6, comprisingremoving at least a portion of the permeate stream through the hydrogenpermeable membrane.
 8. The method of claim 6, wherein the hydrogenpermeable membrane does not include palladium or palladium alloys. 9.The method of claim 6, wherein the hydrogen permeable membrane comprisesa metal doped silica.
 10. The method of claim 1, wherein the catalyticstructure comprises an inner surface layer of the catalytic membranereactor assembly.
 11. The method of claim 1, wherein the metal particlesdo not include a precious metal.
 12. The method of claim 1, wherein themetal particles include aluminum, nickel, and magnesium.
 13. The methodof claim 1, wherein the catalytic structure for the ATR reaction and theWGS reaction is a single layer.